Jet fuel kerosene and gasoline production from gas oils



Nov. 17, 1970 w. l.. JACOBS 3,540,999

JET FUEL KEROSENE AND GASOLINE PRODUCTION FROM GAS OILS Filed aan. 15, 1969 United States Patent O 3,540,999 JET FUEL KEROSENE AND GASOLINE PRODUCTION FROM GAS OILS William L. Jacobs, Crystal Lake, Ill., assignor to Universal Oil Products Company, Des Plaines, Ill., a corporation of Delaware Filed Jan. 15, 1969, Ser. No. 791,229 Int. Cl. Cg 23/00 U.S. Cl. 208-59 7 Claims ABSTRACT OF THE DISCLOSURE A process for converting heavier hydrocarbonaceous material into jet fuel kerosene, and gasoline fractions. The simultaneous production of both jet fuel and gasoline fractions, in maximumA quantities, is afforded through the utilization of a modified series-now system. A two-stage process in which the jet fuel kerosene fraction is produced in the first stage, with the gasoline fraction being produced in the second stage.

APPLICABILITY OF THE INVENTION The present multiple-stage process is directed toward the production of jet fuel kerosene fractionsnfrorn a sulfurous, higher-boiling charge stock, lwhile simultaneously maximizing the yield of gasoline boiling range fractions. The jet fuel kerosene fractions, withdrawn as product streams, seldom require further treatment in order to conform to the current specications imposed upon various sonic, or standard jet fuels. Suitable charge stocks, for utilization as the fresh feed in the present process, are those containing substantial quantities of hydrocarbons having normal boiling points above a temperature of about 550 F., which temperature is generally considered to be the maximum end boiling point of jet fuel kerosene fractions. Therefore, the most common charge stocks will be vacuum gas oils and/or coker gas oils. It is understood, however, that gas oils resulting from a prior conversion process are also well-suited for use herein. For example, those vacuum gas oils which are derived from the conversion of extremely heavy hydrocarbonaceous material are commonly referred in the art as black oils. Such feed stocks are those which are derived from the conversion of a black oil charge stock containing at least about 10.0% by volume of nondistillable hydrocarbons-ie. hydrocarbonaceous material boiling above a temperature of about 1050 F.

The hydrocarbonaceous material contemplated for conversion into jet fuel kerosene and gasoline boiling range fractions, includes a blend of coker gas oil, diesel oil, and light gas oil, which blend has a gravity of about 27.4 API, a `sulfur concentration of about 1.36% by weight, an initial boiling point of about 401 F. and an end boiling point of about 866 F. Other suitable gas oils are: Lloydminster heavy gas oil, containing about 2.2% by weight of sulfur, 600 p.p.m. of nitrogen, and having a gravity of about 21.5 API, an initial boiling point of about 610 F. and an end boiling point of about 810 F.; a Wainwright heavy gas oil having a gravity of 23.2 API, a sulfur concentration of about 1.23% by weight, containing 600 p.p.m. of nitrogen, and having an initial boiling point of about 635 F. and an end boiling point of about 862 F.; a Redwater heavy gas oil having an initial boiling point of about 635 F., an end boiling point of about 855 F., a gravity of about 28.1 API, and containing 0.6% by weight of sulfur and 700 p.p.m. of nitrogen; and, a virgin vacuum gas oil derived from 3,540,999 Patented Nov. 17, 1970 rice a sour Wyoming crude oil, having a gravity of about 21.2 API, an initial boiling point of about 590 F., and an end boiling point of about l050 F., containing 2.42% by weight of sulfur and about 1300 p.p.m. of nitrogen.

For the purposes of denition, gasoline boiling range fractions are herein considered to be that portion of the normally liquid product effluent boiling below the initial boiling point selected for the desired kerosene fraction. Kerosene fractions are generally defined in petroleum technology as having an initial boiling point as low as about 300 F., and an end boiling point of from about 500 F. to about 600 F. Typical jet fuel kerosene fractions have boiling ranges of from about 300 F. to about 550 F., from about 350 F. to about 550 F., from 330 F. to about 500 F., from 375 F. to about 530 F., from 330 F. to about 530 F., from 300 F. to about 530 F., etc. With respect to the motor fuel utilized in internal combustion engines, being gasoline boiling range hydrocarbons, the critical properties are generally considered to be a high octane rating, a particularly specified volatility generally depending upon the locale in which the motor fuel is used, a low degree of olenieity, and low concentrations of contaminating influences. In contrast, many more criteria are employed in describing the fuel for use in jet engines, and even these are further restricted depending upon engine complexity, speed, cruisingl altitude, distance, etc. Specifications for jet fuels of particular physical and/or chemical characteristics are designated as JP-1, JP-3, JP-4, JP-5, JP-6, JP-8, Jet-A and Jet-Al, etc. Although the allowable limits of the individual specifications, employed as criteria in determining a particular grade of jet fuel, may differ one from the other, the selected characteristics are generally the same for all types of jet fuels. These include, for example, the gravity in API, particular volumetric distillation temperatures, the freeze point, the flash point, thermal stability, luminoscity number, the IPT Smoke Point, the aniline point, and, the concentration of contaminating inliuences, particularly sulfur. The abovedesignated jet fuels are described in the literature as standard, or sonic jet fuels. It has already been anticipated that, although a jet fuel kerosene fraction may conform to the specifications for a given standard jet fuel, the same will not be considered suitable for use in future `supersonic jet transports. It is contemplated that the jet fuel specifications, necessitated by supersonic jet transport-s, will be further restricted.

Detailed requirements for various jet fuels may be found in the ASTM Specifications for Aviation Tur-bine Fuels, ASTM designation D-1655-67T. Of these requirements, the three most critical are considered to be the IPT Smoke Point, generally not less than 25 mm., the concentration of aromatic hydrocarbons, generally less than about 20.0% by volume and the concentration of sulfur.

The demand for standard jet fuel of the foregoing types, by the airline industry in the United States, has been estimated to increase to about 1.12 million barrels per day in the year 1975. The potential supply, of petroleum stocks within the boiling range of say about 300 F. to about 550 F., is estimated at about 2.20 million barrels per day in 1975, which appears to be more than ample to satisfy the estimated demand. It must, however, be recognized that this particular boiling range fraction also contributes to the supply of other fuels including .TP-4 and IP-S military jet fuels, diesel fuels, heating oils and kerosene, gasoline and petrochemical feed stocks.

Prior to 1975, the advent of jumbo jets and supersonic transports will increase the demand for jet fuel kerosene fractions by virtue of significantly increased fuel consumption brought about by increased size and speed. It

should further be noted that the lower boiling portion of the previously described jet fuel fractions will deplete the supply of gasoline boiling range hydrocarbons. That is, from about 10.0 vol. percent to about 50.0 vol. percent of most jet fuel kerosene fractions consist of motor fuel components boiling below about 400 F., which temperature is generally considered the maximum end boiling point for gasoline fractions. While it appears certain that the overall effect of the future jumbo jets and supersonic jet transports will be an increase in the demand for greater quantities of jet fuels, it is not possible to project the exact quantity in terms of millions of barrels per day. Likewise, no reasonable estimate can be made with respect to the effect on the demand for -motor fuel gasolines. As the supersonic jet transports displace the standard jets in use today, the demand for standard jet fuel will certainly decreases, although the combined demand for sonic and supersonic jet fuels will certainly increase. This complex situation is further complicated by the fact that automotive transportation is steadily increasing, placing a corresponding increase on the demand for motor fuel gasoline fractions. The process encompassed by my invention recognizes that there will be a critical need for both jet fuel kerosene and motor fuel gasoline fractions, and affords the simultaneous production of both in an economical manner.

OBJECTS AND EMBODIMENTS An object of my invention is to provide a process for converting heavier hydrocarbonaceous material into jet fuel kerosene fractions, accompanied by maximum production of normally liquid gasoline boiling range hydrocarbons.

Another object is to produce jet fuel kerosene fractions meeting the requirements imposed upon the IPT Smoke Point, aromatic concentration and sulfur content.

Another object is to provide a semi-series flow, multiple-reaction zone process for the production of jet fuel kerosene and gasoline fractions.

Therefore, in one of its embodiments, the present invention affords a process for producing gasoline and jet fuel kerosene fractions from a sulfurous, heavier hydrocarbon charge stock, which process comprises the steps of: (a) reacting said charge stock 'with hydrogen in a first catalytic reaction zone, at a maximum catalyst bed temperature below about 850 F. and a pressure greater than about 1000 p.s.i.g., said temperature and pressure selected to convert sulfurous compounds to hydrogen sulfide and hydrocarbons; (b) separating the resulting first reaction zone effluent in a first separation zone, at substantially the same pressure and at a temperature of from about 550 F. to about 750 F., to provide a first vaporous phase and a first liquid phase; (c) further separating said first vaporous phase in a second separation zone, at substantially the same pressure and at a lower temperature in the range of from about 60 F. to about 140 F., to provide a hydrogen-rich second vaporous phase and a second liquid phase; (d) reacting said first liquid phase with hydrogen in a second catalytic reaction zone containing a siliceous catalytic composite comprising a Group VIII metallic component, at a maximum catalyst bed temperature below about 800 F. and a pressure greater than about 1000 p.s.i.g.; (e) separating the resulting second reaction zone effluent in a third separation zone, at substantially the same pressure and at a temperature in the range of from about 60 F to about 140 F., to provide a hydrogen-rich third vaporous phase and a third liquid phase; and, (f) introducing said third liquid phase and said second liquid phase into a fourth separation zone, at conditions of temperature and pressure to provide a 'fourth liquid phase having an initial boiling point of from 500 F. to about 600 F., and recovering a gasoline and a jet fuel kerosene fraction.

Other embodiments of my invention are directed towards preferred processing techniques, operating condi- CII tions and the various catalytic composites for utilization in the multiple reaction zones. For example, one such technique, preferred when a crystalline aluminosilicate catalyst is employed, involves introducing the second reaction zone eiuent into the same cold separator as the rst vaporous phase from the hot separator. This eliminates one cold separator, a transfer line, and permits the use of common recycle facilities to both reaction zones. Another operational technique involves reacting the fourth liquid phase in said second catalytic reaction zone. The preferred catalytic composite, for utilization in the first reaction zone is a siliceous composite containing at least one metallic component from Group VI-B and VIII of the Periodic Table, While the catalytic composite disposed in said second catalytic reaction zone is preferably a crystalline aluminosilicate containing a Group VIII metallic component selected from the group consisting of nickel, platinum and palladium. These, as Well as other objects and embodiments of my invention, will be evident from the following, more detailed description thereof.

SUMMARY OF INVENTION As hereinbefore set forth, the primary purpose of my invention is to provide a process which affords the simultaneous production of jet fuel kerosene and gasoline boiling range fractions. A corollary purpose is to provide a process having the built-in flexibility necessary to adjust the respective proportions of the two principal product streams to satisfy an ever-changing, unpredictable marketing demand. Thus, in accordance with my invention, a given charge stock can be processed in such a manner that the jet fuel kerosene yield is 20.0 vol. percent, based upon a fresh feed, or the jet fuel kerosene fraction produced is about 55.0 vol. percent, based upon fresh feed. The operation of the process is facilitated when effected to produce a given yield of the jet fuel kerosene fraction, While maximizing the production of motor fuel gasoline components from the remainder of the fresh charge stock. Briefly, in the `first reaction zone, the catalytic composite and operating conditions are selected to assure virtually complete desulfurization of the fresh gas oil charge stock, while simultaneously converting heavier hydrocarbons into the jet fuel kerosene fraction. The product effluent from the first reaction zone is separated in a manner such that the heavier portion thereof serves as a portion of the charge, in semi-series flow, to the second reaction zone. A principal function of the second catalytic reaction zone is to crack the heavy gas oil components into maximum quantities of gasoline boiling range components.

Before further summarizing my invention, it is believed that several other definitions are necessary in order that a clear understanding be made available. In the present specification, as well as the appended claims, the use of the term pressure substantially the same as is intended t0 connote that the pressure imposed upon a given vessel, or section of the system, is the same as the pressure imposed upon the vessel immediately upstream therefrom, allowing only for the pressure drop naturally occurring as a result of the flow of uids through the system. Similarly, the term temperature substantially the same as is intended to indicate that the temperature of a given stream entering a particular vessel is substantially the Same as the temperature of the stream as it emanates from the vessel immediately upstream therefrom. Taken into consideration is, of course, the temperature drop due to heat loss, and to that from the conversion of sensible to latent heat due to vaporization. Where utilized herein, the term semi-series flow alludes to the fact that the portion of the product eilluent from the first catalytic reaction zone utilized as the charge to the second catalytic reaction zone, is introduced into the latter at substantially the same pressure. In many instances, this charge to the second reaction zone need not be heated but can be introduced at substantially the same temperature it has as it emanates from the first separation zone.

The present invention involves the utilization of a first catalytic reaction zone having disposed therein a catalyst comprising at least one metallic component from Group VI-B and the Iron-Group of the Periodic Table, the primary purpose of which is to effect substantially complete desulfurization of the hydrocarbonaceous feed stock at comparatively rnild severities. The catalytic cornposite, utilized in the second reaction zone for the purpose of hydrocracking the gas oil boiling range material into motor fuel gasoline components, comprises a Group VIII metallic component selected from the group consisting of nickel, platinum and palladium. Intermediate the two zones is a hot separator, functioning at substantially the same pressure as the first reaction zone and at a temperature in the range of from about 500 F. to about 800 F., and preferably from about 550 F. to about 750 F. The liquid phase from the hot separator, at substantially the same pressure, is introduced into the second catalytic reaction zone, often without substantial change in temperature.

The catalytic composites utilized in the present process comprise metallic components selected from the metals of the Group VI-B and VIII of the Periodic Table, and compounds thereof. Thus, in accordance with the Periodic Table of The Elements, E. H. Sargent & Co., 1964, suitable metallic components are those selected from the group consisting of chromium, molybdenum, tungsten, iron, ruthenium, osmium, cobalt, rhodium, iridium, nickel, palladium and platinum. While neither the precise composition, nor the method of manufacturing the various catalytic composites is considered essential to my invention, certain aspects are preferred. For example, since the charge stocks to the present process are gas oil fractions, a considerable portion of which boils above the kerosene boiling rangei.e. above about 550 F.-it is preferred that the components of the catalyst possess the propensity for effecting hydrocracking while simultaneously producing a substantially sulfur-free normally liquid hydrocarbon product. Suitable catalytic composites, for use in effecting the desulfurization reactions in the first reaction zone, generally comprise from about 4.0% to about 40.0% by weight of a Group VI-B metallic component, and from about 1.0% to about 6.0% by weight of an Iron-Group metallic component. It is understood that these concentrations, as well as those hereinafter set forth, are computed on the basis of the elemental metals, regardless of the precise state in which they exist within the catalytic composite. These catalytically active metallic components are generally composited with a suitable siliceous refractory inorganic oxide carrier material, the quantity of silica determining the degree of hydrocracking activity. Suitable refractory inorganic oxides include alumina, zirconia, magnesia, titania, thoria, boria, hafnia, etc., and mixtures thereof. Another group of suitable carrier materials are those having combined therewith from about to about 35.0% by weight of boron phosphate. In general, the silica to alumina weight ratio will be Within the range of from about 90 to about 80/20.

With respect to the second catalytic reaction zone, the catalyst comprises metallic components from Group VIII in amounts within the range of from about 0.01% to about 20.0% by weight. These catalytic components are also combined with one or more of the foregoing refractory inorganic oxides. In a particularly preferred embodiment, the catalyst carrier material is a crystalline aluminosilicate, or zeolitic material.

As hereinbefore set forth, the preferred catalytic composite for use in the second catalytic reaction zone comprises a crystalline aluminosilicate, commonly referred to in the art as a molecular sieve, or zeolitic material. Molecular sieves are characterized by a crystalline structure having many small cavities connected by smaller pores of uniform size. These pores may vary in size from 3 angstrom units up to about 15 angstrom units. EX- emplary of such crystalline aluminosilicates are Type A, Type X, Type Y, Type U, faujasite, etc. In accordance with the present process, maximum gasoline boiling range hydrocarbons are realized through the utilization of a pilled, binderless faujasite having pore openings in the vicinity of about l0 angstroms, and which is further characterized by being at least 90.0% by weight zeolitic material, as distinguished from amorphous material. A Group VIII metallic component, selected from the group consisting of platinum, palladium, nickel and compounds thereof, is added to the faujasite particles in amounts of from about 0.01% to about 20.0% by weight. Lesser quantities of the noble metals, platinum and palladium, up to about 3.0% by weight are generally preferred. These catalytic components, various compositions thereof, molecular sieves, refractory inorganic oxides, and the methods of the manufacture thereof are well described in the prior art, and additional discussion herein is not believed necessary.

In practicing the present invention, the charge stock is admixed with recycled hydrogen in an amount of about 3000 to about 20,000 standard cubic feet per barrel. Following suitable heat-exchange with various hot product effluent streams, the hydrocarbon/hydrogen mixture is heated to a temperature level such that the catalyst bed temperature is controlled within he range of about 600 F. to a maximum of 850 F. The catalyst bed inlet temperature is regulated to control the outlet temperature at a maximum level of 850 F., and preferably not higher than 800 F. Since the principal reactions are exothermic in nature, a temperature rise will be experienced as the charge stock passes through the catalyst bed. A preferred technique limits the temperature in this first catalytic reaction zone to about F., and the use of conventional quench streams, at one or more intermediate loci of the reaction zone, is contemplated for this purpose. The reaction zone contains, for example, a catalyst of 1.8% by weight of nickel and 16.0% by weight of molybdenum combined with a carrier material of 63.0% by weight of alumina and 37.0% by 'weight of silica. The reaction zone is maintained under an imposed pressure of from about 1000 to about 4000 p.s.i.g., and the liquid hourly space velocity (defined as volumes of liquid hydrocarbon charge per hour per volume of catalyst) is in the range of from about 0.4 to about 3.5.

The total product efiiuent, without substantial change in pressure, although sometimes at a lower temperature of from about 550 F. to about 750 F., as a result of the use of the product effluent as a heat-exchange medium, is introduced into a hot separator. A vaporous phase, comprising hydrogen, hydrogen sulfide, ammonia, normally gaseous hydrocarbons, butanes, pentanes and heavier hydrocarbons boiling below a temperature of about 550 F., is withdrawn from the hot separator and introduced, at a temperature in the range of from about 60 F. to about F., into a cold separator. The temperature of the stream entering the hot separator is controlled at a level which insures that substantially all of the jet fuel kerosene components, for example, boiling from about 300 F. to about 550 F., are carried over in the vapor phase, while substantially all of the hydrocarbon- `aceous material boiling above a temperature of about 550 F., is withdrawn from the hot separator as alliquid phase. Condensed, normally liquid hydrocarbons are removed from the cold separator, and subjected to a product recovery systemi.e. a fractionating column-in order to recover the desired jet fuel fraction. In one embodiment, primarily dependent upon the desired product distribution between the motor fuel gasoline and jet fuel kerosene fractions, a portion of the liquid phase withdrawn from the hot separator is recycled to combine with the fresh gas oil charge stock to the first reaction zone. Since the heavier components, boiling above 550 F., are once again subjected to the environment conducive in part to hydrocracking, this technique affords fiexibility with respect to overall product distribution. Suitable combined feed ratios, defined as volumes of total liquid charge per volume of fresh gas oil charge, are within the range of from about 1.1 to about 4.5. That portion of the liquid phase not being recycled from the hot separator to combine with the fresh gas oil charge serves in part as the charge to the second catalytic reaction zone. As hereinbefore stated, the liquid phase from the cold separator is subjected to a product recovery system in order to recover the jet fuel kerosene fraction. A bottoms fraction, containing those hydrocarbons boiling above the desired end boiling point of the jet fuel fraction, is combined with the liquid phase from the hot separator as charge to the second catalytic reaction zone. In those instances where the quantity of jet fuel kerosene, produced within the first catalytic reaction zone, exceeds the amount intended to be produced, provision is made for combining the excess with the heavy bottom stream and the liquid phase from the hot separator. The catalyst disposed within the second reaction zone is a piled faujasite crystalline aluminosilicate, of which about 92.5% by weight is zeolitic, combined with 5.0% by weight of nickel. The reaction zone is maintained at a pressure above about 1000 p.s.i.g., a practical upper limit being about 3000 p.s.i.g. The hydrogen circulation rate is at least about 3000 standard cubic feet per barrel, with an upper limit of about 15,000 standard cubic feet per barrel, based upon the total liquid feed, including the excess jet fuel kerosene produced in the first reaction zone. The liquid hourly space velocity, previously defined herein, is within the range of from about 0.5 to about 4.0. It is particularly preferred to maintain the maximum catalyst bed temperature at a level not exceeding 800 F. In many applications of the present invention, the inlet temperature of the catalyst bed within the second catalytic reaction zone will be in the range of from about 550 F. to about 700 F. Thus, as hereinbefore set forth, there will be many instances where the charge to the second reaction zone will not have to be increased in temperature. As with the first catalytic reaction zone, a temperature rise is experienced as the charge passes through the catalyst bed. It is preferred, with respect to the second catalytic reaction zone, to limit such temperature rise to about 50 F., particularly when the catalytically active metallic component is either platinum, or palladium, and the use of conventional quench streams is again contemplated for this purpose. The total product effluent from the second catalytic reaction zone is introduced into a second cold separator, at substantially the same pressure and at a temperature in the range of from about 60 F. to about 140 F. The principally vaporous phase from the second cold separator is substantially clean with respect to hydrogen sulfide in view of the fact that complete desulfurization has been effected in the first reaction zone. Thus, there is no requirement to treat the recycled hydrogen stream in H25 removal facilities. This particular scheme is preferred in those instances where the hydrocracking catalyst in the second reaction zone is sulfursensitive. However, in those instances where the second zone catalyst is a crystalline aluminosilicate, or zeolitic in nature, hydrogen sulfide can be tolerated in the second zone. In these cases, the product effluents from both reaction zones are introduced in to a common cold separator, and common recycle facilities can be utilized. When two cold separators are utilizedeg the second zone catalyst is sulfur-sensitive-the operation of the overall process is facilitated, particularly with respect to commerciallyscaled units, when the vapor phases from both cold separators are combined and introduced into a hydrogen sulfide removal system.

The normally liquid hydrocarbon portion of the product effluent emanating from the second catalytic reac tion zone, is recovered from the cold separator and introduced into product separation means. The process is facilitated by employing the same fractionation system utilized in separating the normally liquid hydrocarbon stream from the first cold separator. Typical of the component separations effected with respect to the total normally liquid product effiuent is a normally gaseous stream comprising butanes, lighter hydrocarbons and other gaseous components; a pentane/ hexane fraction which may be utilized as a motor fuel blending component; and, a heptane to 350 F. motor fuel fraction, or other gasoline fractions boiling below the desired initial boiling point of the jet fuel kerosene. This motor fuel fraction may be utilized in combination with other similarly constituted refinery streams as charge stock to a catalytic reforming system. The desired jet fuel kerosene, for example, having a boiling range from 350 F. to 550 F. is also recovered as a product stream. A bottoms fraction, containing the hydrocarbonaceous material boiling at a temperature above the desired end boiling point of the jet fuel kerosene fraction is recovered and combined with the hot separator liquid phase as the charge to the second catalytic reaction zone. In the event that the quantity of jet fuel kerosene exceeds the predetermined, desired amount, the excess is conveniently admixed with the charge to the second catalytic reaction zone, therein converted into gasoline boiling range hydrocarbons.

DESCRIPTION OF DRAWINGS The process encompassed by my `invention is more clearly understood by reference to the accompanying drawing which illustrates one embodiment thereof. In the drawing, only those vessels and process lines required for an understanding of the embodiment have been included. Miscellaneous appurtenances, including valves, pressurereducing valves, controls, instruments, pumps, compressors, heat-exchangers, start-up lines and heat-recovery circuits have either been reduced in number or completely eliminated. The use of this kind of conventional hardware is Well within the purview of those skilled in the techniques of petroleum refining processing. It is further understood that the drawing is presented for the sole purpose of illustration, and is not intended to be limited to the particular charge stock, quantities, rates, operation conditions, product yield and/or distribution, etc., employed by way of illustration. With reference now to the drawing, the charge stock, for example a heavy Vacuum gas oil from a Lloydminster crude, the property inspections of which are presented in the following Table I, is introduced into the process by Way of line 1.

TABLE I.-CHARGE STOCK PROPERTIES Gravity, API 21.5 ASTM distillation, F.:

IBP 610 10% 660 30% 690 50% to 70% 710 to 730 765 End point 810 Sulfur, wt. percent 2.20 Nitrogen, p.p.m. 600

The charge stock continues through line 1, being admixed with a recycled hydrogen stream from line 2, in an amount of about 6000 standard cubic feet per barrel. The charge stock rate is 10,000 barrels per day, and it is intended that this charge stock be converted into a jet fuel kerosene, in a yield of about 20.0% by volume, based upon fresh feed, having a boiling range of about 350 F. to about 550 F., and maximum quantities of a heptane to 350 F. motor fuel gasoline fraction. In the indicated embodiment, the charge stock/hydrogen mixture following heat-exchange with relatively hot product efuent streams, continues through line 1 into heater 3. With this particular charge stock and the desired distribution of the product,

the operational techniques do not call for a liquid recycle stream to be admixed with the charge stock in line 1. When this technique is deemed advisable, the recycle material is admixed with the charge stock to provide a combined liquid feed ratio of from about 1.1 to about 4.5.

Heater 3 raises the temperature of the incoming hydrogen/charge stock stream to a level of about 675 F., as measured at the inlet to the catalyst bed disposed in reactor 5. The thus-heated mixture is introduced downflow by way of line 4, and is withdrawn from reactor 5 by way of line 6 at a temperature of about 775 F. The catalyst disposed within reactor 5 is a desulfurization catalyst having some hydrocracking activity, and comprises 1.8% by weight of nickel and 16.0% by weight of molybdenum, computed as the elemental metals, combined with a carrier material of 63.0% by weight of alumina and 37.0% by Weight of silica. The charge stock, at a pressure of about 2000 p.s.i.g., passes through the catalytic composite at a liquid hourly space velocity of 1.01.

Component analyses of the product eiiluent withdrawn from reactor 5 by way of line 6, to be introduced thereby into hot separator 7, are presented in the following Table II. The yields indicated therein reflect a hydrogen consumption (exclusive of solution loss) of about 1208 standard cubic feet per barrel, or 1.97% by weight, 'based upon the fresh gas oil charge stock.

TABLE II.-REACTR 5 PRODUCT EFFLUENT Component Wt. percent Vol. percent Prior to entering hot separator 7, the hot reaction zone effluent in line 6 is utilized as a heat-exchange medium to lower its temperature to a level of about 600 F., and thereafter continues through line 6. A rst principally vaporous phase, containing substantially all of the material boilingr below a temperature of about 550 F., and substantially free from hydrocarbonaceous material boiling above about 550 F., is removed from hot separator 7 by Way of line 8, cooled to a temperature of about 120 F., and introduced into cold separator 9. A principally vaporous phase is withdrawn from cold separator 9, by way of compressive means not illustrated in the drawing, and continues through line 2 to be admixed with the charge stock in line 1. Make-up hydrogen, to compensate for that consumed within the process, may 'be introduced from any suitable external source, at any suitable location in the process system. Generally, such make-up hydrogen is introduced on the upstream, or suction side of the compressive means. As hereinbefore set forth, the vaporous phase from line 2 may be treated, when necessary, for the removal of gaseous constituents other than hydrogen in order that the hydrogen concentration in the recycled stream is about at least 80.0 mol percent. Such treating facilities are well known in the prior art, and, therefore, are not indicated in the illustrated embodiment.

A principally liquid phase is withdrawn from cold separator 9 by way of line 10, and is introduced thereby into 10 fractionator 11. This stream contains all of the hydrocarbonaceous material boiling at temperatures of 550 F. and below, and a relatively minor quantity of hydrocarbons boiling above a temperature of 550 F. Fractionator 11 functions at conditions of temperature and pressure which permits the recovery of the jet fuel kerosene fraction boiling from 350 F. to 550 F.; this product is indicated as being withdrawn from fractionator 11 by way of line 21. The excess kerosene fraction, in the present illustration 6.58 vol. percent 'based upon fresh feed, is diverted from line 21 through line 22 containing valve 23. An overhead stream comprising butanes, lighter normally gaseous hydrocarbons and other gaseous material is withdrawn by way of line 18 and may be further separated in order to recover particularly desired components. A pentane/hexane fraction is withdrawn through line 19, and may be utilized, at least in part, in a pool as motor fuel blending components. The gasoline fraction, comprising heptanes and heavier hydrocarbonaceous material boiling up to a temperature of about 350 F., is

withdrawn by way of line 20. Material boiling above the desired end point of the standard jet fuel fraction, 550 F., is withdrawn by way of line 24.

The principally liquid phase withdrawn from hot separator 7 by way of line 12, contains substantially all the hydrocarbonaceous material boiling above a temperature of 550 F., and is introduced by way of line 12 into reactor 13 at a catalyst bed inlet temperature of about 675 F. Prior to being introduced into reactor 13, the liquid phase is admixed with the 550 F-plus material emanating from fractionator 11 as a bottoms stream in line 24, and the excess kerosene fraction in an amount of about 650 barrels per day from line 22. The liquid mixture is further admixed with about 9000 standard cubic feet per barrel of hydrogen by way of line 16. The catalyst disposed within reactor 13 is a composite of binderless faujasite, about 91.7% by weight of which is zeolitic, and 5.0% by weight of nickel, calculated as the elemental metal. Reactor 13 is maintained at a pressure of about 2000 p.s.i.g. and the liquid hourly space velocity, based lupon 7846 barrels per day of fresh feed, is 0.65. The product effluent from reactor 13 is withdrawn by way of line 14, cooled to a temperature of about F. and introduced into cold separator 15. A principally vaporous phase is withdrawn from cold separator 15 by way of line 16, by compressive means not illustrated in the drawing, to combine with the fresh feed in line 12. A principally liquid phase, consistnig primarily of normally liquid hydrocarbons, is withdrawn from cold separator 15 by way of line 17, and introduced therethrough into fractionator 11. As hereinbefore stated, when the second zone catalyst is zeolitic in nature, as in the present illustration, and hydrogen sulfide can be tolerated in the reaction mixture, the second zone product effluent in line 14 can be introduced into cold separator 9 along with the vaporous phase in line 8. Those having skill in the art of petroleum refining processes and technique, will be cognizant of the design of the common recycle facilities, and further discussion herein is not required.

Although the liquid phase in line 17 is indicated as being admixed with the liquid phase in line 10, prior to the introducton thereof into fractionator 11, it is understood that the two streams may utilize separate and individual feed points. The utilization of a single fractionation system as indicated is preferred from the standpoint of facilitating the separation of the total reaction zone product into the desired product stream.

Component stream analyses of the product effluent from reactor 13 are presented in the following Table III. The values indicated in Table III reflect a hydrogen consumption of about 842 standard cubic feet per barrel, or 1.37 wt. percent. The 6,277 barrels per day of 550 F.-plus material, in the liquid phase in line 17, are re-introduced as recycle to reactor 13 by way of the bottom stream in line 24.

TABLE IIL-REACIOR 13 PRODUCT EFFLUENT The overall product distribution and liquid yields, based upon the fresh gas oil charge rate of 10,000 barrels per day, is presented in the following Table IV.

TABLE IV.OVERALL LIQUID YIELDS AND DISTRIB UTION Component Wt. percent Vol. percent Bbl/day Butanes 11. 54 1-8. 72 1, 872 22. 56 31. 73 3. 173 Heptane-350 F 45. 65 55. 08 5, 508 3 F.-55 18. 05 20.00 2, 000

The foregoing Table IV indicates that, from the 10,000 barrels of fresh gas oil charge, 5,508 barrels per day of a gasoline fraction were simultaneously produced with the intended quantity of 2,000 barrels per day of the jet fuel kerosene fraction. It should further be noted that the total quantity of butanes, pentanes, and hexanes, well suited for utilization in a motor fuel blending pool, amounted to 5,045 barrels per day. With respect to the pentane/hexane fraction, the gravity is 83.6 API and the F-1 clear octane rating is 85.7 (increased to 98.6 upon the addition of 3 ml. TEL). The heptane to 350 F. gasoline fraction has a gravity of 53.1 API, contains 35.6 vol. percent parains, 53.6 vol. percent naphthenes and 10.7 vol. percent aromatics, and constitutes an excellent charge to a catalytic reforming unit. The jet fuel kerosene fraction, as produced and without further treatment, indicates less than about 1.0 p.p.m. of sulfur, an IPT Smoke Point of 25 mm. and contains 13.0 Vol. percent aromatics.

EXAMPLE This example is presented to illustrate the results obtained through the utilization of my invention in a situation where the desired quantity of the jet fuel kerosene fraction is significantly increased. This example is based upon a commercially-scaled unit having a design capacity of about 40,000 barrels per day; in the illustration, the charge rate is 36,000 barrels per day. The charge stock is a mixture of light atmospheric gas oil, light vacuum gas oil, and heavy vacuum gas oil derived from a full boiling range California crude. Pertinent properties of the gas oil charge stock are presented in the following Table V.

TABLE V.-'GAS OIL BLEND PROPERTIES It is intended that this gas oil charge stock be converted to the extent of producing about 62.0 vol. percent of a .IP-8 jet fuel kerosene fraction having a boiling range of from about 300 F. to about 550 F., while maximizing the production of gasoline boiling range material from the remainder of the charge stock. The operation is effected in a two-stage reaction zone system of the type previously described with respect to the embodiment illustrated in the accompanying drawing. The first catalytic reaction zone is maintained at a pressure of about 2,000 p.s.i.g., and a catalyst bed inlet temperature of 750 F. The liquid hourly space velocity is 0.60 and the hydrogen circulation rate is 6,000 standard cubic feet per barrel. The catalyst disposed within the first reaction zone is a composite of 1.8% by weight of nickel and 16.0% by weight of molybdenum, computed as the elemental metals, combined with a carrier material of 63.0% by weight of alumina and 37.0% by weight of silica.

The second catalytic reaction zone is maintained at a pressure of 2,000 p.s.i.g. and a catalyst bed inlet temperature of about 700 F. The hydrogen circulation rate is 8,000 standard cubic feet per barrel, and the liquid hourly space velocity is about 2.0. The catalyst disposed within the second catalytic reaction zone is a composite of 75.0% by weight of silica and 25.0% by weight of alumina, with which is combined 0.4% by Weight of plantinum, calculated as the element.

The hydrogen consumption in the first catalytic reaction zone is 1,345 standard cubic feet per barrel, or 2.18% by weight; in the second catalytic reaction zone, the hydrogen consumption is 1.36% by weight, or 740 standard cubic feet per barrel, for a total of 2,'085 standard cubic feet per barrel, or 3.54% by weight. These figures are reflected in the following Tables VI, VII and VIII. Table VI indicates the component analyses of the product effluent from the first catalytic reaction zone, Table VII the product efiluent distribution from the second reaction zone, and Table VII indicates the total liquid yields and distribution thereof.

Component Wt. percent Vol. percent 11 1. 7 Pentanes- 0. 1. 3g Hexane 1. 11 1. 49 Heptane-300 F. 2. 02 2. 47 300 F550 F 28. 78 31. 72 550 F.-plus 66. 04 71. 20

TABLE VIL-SECOND STAGE PRODUCT EFFLUENT Component Wt. percent Vol. percent Ammonia Hydrogen sullde Methane Ethane. 0. 03 Propane. 1. 29 Butanes- 6. 30 Pentanes 6. 63 Hexane 5. 76 Heptane-300 F 17. 96 300 F.-550 F 29. 61

TABLE VKL-OVE RALL LIQUID YIELDS AND DIST RIB UTION Component Wt. percent -Vol. percent Bbl/day Butanes 7. 42 12. 10 4, 430 Pentanes/hexanes 14. 22 20. 14 7, 360 Romane-300 F 19. 98 24. 40 8, 975 300 IPs-550 F 58. 39 66.00 24, 100

The pentane/ hexane fraction indicates .a gravityof 83.7 API, and an F-l clear octane rating of 84.0 (increased to 98.2 with an addition to 3 ml. TEL). The heptane to 300 F. gasoline fraction has a gravity of 55.0, and contains 33.2 vol. percent parafns, 61.8 vol. percent naphthalenes and 5.0 Vol. percent aromatics. The 300 F. to 550 F. jet fuel kerosene fraction has a gravity of 40.0 API, a flash point of 110 F., an IPTVSmoke Point of 25 mm., and contains less than about 1.0 p.p.m. of sulfur and 8.0 vol. percent aromatic hydrocarbons.

The foregoing specification, and particularly the examples, clearly indicates the method of effecting the pres- 13 ent invention and the benefits aorded through the utilization thereof.

I claim as my invention:

1. A process for producing gasoline and jet fuel kerosene fractions from a sulfurious heavier hydrocarbon charge stock which comprises the steps of:

(a) reacting said charge Stock With hydrogen in a first catalytic reatcion zone at a maximum catalyst bed temperature below about 850 F. and a pressure greater than about 1000 p.s.i.g., said temperature and pressure selected to convert sulfurous compounds to hydrogen sulfide and hydrocarbons;

(b) separating the resulting first reaction zone eliiuent in a lirst separation zone, at substantially the same pressure and at a temperature of from about 550 F to about 750 F., to provide a first vaporous phase and a first liquid phase;

(c) further separating said first vaporous phase in a second separation zone, at substantially the same pressure and at a lower temperature in the range of from about 60 F. to about 140 F., to provide a hydrogen-rich second vaporous phase and a second liquid phase;

(d) reacting said first liquid phase with hydrogen in a second catalytic reaction zone containing a siliceous catalytic composite comprising a Group VIII metallic component, at a maximum catalyst bed temperature below about 800 F. and a pressure greater than about 1000 p.s.i.g.;

(e) separating the resulting second reaction zone efliuent in a third separation zone, at substantially the same pressure and at a temperature in the range of from about 60 F. to about 140 F., to provide a hydrogen-rich third vaporous phase and a third liquid phase; and,

(f introducing said third liquid phase and said second liquid phase into a fractional distillation zone, at conditions of temperature and pressure to provide a fourth liquid phase having an initial boiling point of from 500 F. to about 600 F., and recovering a gasoline and a jet fuel kerosene fraction.

2. The process of claim 1 further characterized in that said fourth liquid phase is reacted with hydrogen in said second catalytic reaction zone.

3. The process of claim 1 further characterized in that at least a portion of said second and third vaporous phases are recycled to said first and second catalytic reaction zones respectively.

4. The process of claim 1 further characterized in that the catalyst disposed in said iirst reaction zone is a siliceous composite containing at least one metallic component from Groups VI-B and VIII of the Periodic Table.

5. The process of claim 1 further characterized in that the catalytic composite disposed in said second catalytic reaction zone is a crystalline aluminosilicate containing a Group VIII metallic component selected from the group consisting of nickel, platinum and palladium.

6. A process for producing gasoline and jet fuel kerosene fractions from a sulfurious heavier hydrocarbon charge stock when comprises the steps of:

(a) reacting said charge stock with hydrogen in a irst catalytic reaction zone at a maximum catalyst bed temperature below about 850 F. and a pressure greater than about 1000 p.s.i.g., said temperature and pressure selected to convert sulfurous compounds to hydrogen sulfide and hydrocarbons;

(b) separating the resulting first reaction zone effluent in a first separation zone, at substantially the same pressure and at a temperature of from about 550 F. to about 750 F., to provide a iirst vaporous phase and a first liquid phase;

(c) further separating said rst vaporous phase and a previously hydrocracked product effluent, in a second separation zone, at substantially the same pressure and at a lower temperature in the range of from about F. to about 140 F., to provide a hydrogen-rich second vaporous phase and a second liquid phase;

(d) reacting said first liquid phase with hydrogen in a second catalytic hydrocracking reaction zone containing a siliceous catalytic composite comprising a Group VIII metallic component and a crystalline aluminosilicate, at a maximum catalyst bed temperature below about 800 F. and a pressure greater than about 1000 p.s.i.g.;

(e) separating the resulting hydrocracked product efiiuent in said second separation zone as aforesaid; and,

(f) introducing said second liquid phase into a fractional distillation Zone, at conditions of temperature and pressure to provide a third liquid phase having an initial boiling point of from 500 F. to about 600 F., and recovering a gasoline and a jet fuel kerosene fraction.

7. The process of claim 6 further characterized in that said third liquid phase is reacted with hydrogen in said second catalytic reaction zone.

References Cited UNITED STATES PATENTS 3,132,087 5/1964 Kelley et al 208-60 DELBERT E. GANTZ, Primary Examiner. A. RIMENS, Assistant Examiner.

U.S. Cl. X.R. 20S- 60, 15, 16 

